Aromatics production process



Aug. 20, 1968 G. E. ADDISON AROMATICS PRODUCTION PROCESS Filed Nov. 2, 1966 a mw\ //V VE/V TOR George E. Add/'son A TTOR/VEYS United States ABSTRACT OF THE DISCLOSURE Process of removing aromatics from a reformer feed by solvent extraction, reforming the raffinate, separately solvent extracting the reformed raiiinate to separate aromatics from a second raffinate phase, and recovering the aromatics and the second raffinate phase as products of the process.

This invention is a continuation-in-part of my copending application Ser. No. 460,894, filed June 3, 1965, now abandoned.

This invention relates to a process 4for the enhanced production of aromatics from a hydrocarbon feed in which a reforming step and two solvent extraction steps are performed. More specifically, this invention relates to the solvent extracting of a hydrocarbon feed to remove aromatics therefrom prior to reforming the feed in a process for the production of aromatics.

In one of its embodiments, this invention relates to a process for the enhanced production of aromatic hydrocarbons from a hydrocarbon feed which comprises the steps: solvent extracting the hydrocarbon feed to produce a first rich solvent extract containing aromatics and a rainate; reforming the raffinate; solvent extracting the reformed raffinate to produce a second rich solvent extract containing aromatics and a raffinate; and separating the aromatics from the rich solvent extracts.

In another of its embodiments this invention relates to a combination extraction and reforming process for the enhanced production of aromatics from a hydrocarbonaceous feed mixture which comprises: contacting the feed mixture containing aromatic hydrocarbons with a lean solvent selective for aromatics in a first solvent extractor and forming therein a rich solvent phase containing aromatics and a raffinate phase containing less absorbed compounds; Vwithdrawing the rafiinate phase from the first extractor and introducing it into a reforming catalyst reactor maintained at reforming conditions; withdrawing a normally liquid reactor eiiiuent from the reforming step and contacting it with a lean solvent selective for aromatics in a second solvent extractor to form a rich solvent phase containing aromatics and a raffinate phase containing less absorbed components; combining the rich solvent phase from the first extractor and the rich solvent phase from the second extractor; removing the dissolved aromatics from the combined rich solvent phase to form an aromatics stream and a lean solvent stream; returning a portion of the lean solvent to the first extractor; and returning the other portion of the lean solvent to the second extractor.

High purity aromatic streams are produced from petroleum fractions in large quantities in many refineries today. Typically, a petroleum fraction of desired boiling range is reformed in a catalytic reactor wherein most of the naphthenes in the feed are dehydrogenated to produce aromatics, a few of the paraflins in the 4feed are dehydrocyclized to produce aromatics and the aromatics in the feed are unchanged. Thus the reforming step produces substantial quantities of aromatic hydrocarbons in high concentrations in the normally liquid product called the atent C 3,398,083 Patented Aug. 20, 1968 reformate. In addition the reformate contains asubstantial amount of paraiins and a smallY amount of naphthenes. The reformate has a high octane number because of its high aromatic content and may be used in high octane gasoline fuel service. However, in the past few years it has become increasingly economical to utilize the aromatics in the reformate for petrochemical uses. Thus benzene is a valuable starting material in the manufacture of a variety of petrochemicals. Likewise, ethylbenzene is utilized in the production of styrene, ortho-xylene is utilized in the production of phthalic acid, para-xylene is employed to produce terephthalic acid, etc. However, in order to use the aromatics for petrochemicals, etc., it is necessary to recover them in substantially pure form. Re-

r formates may typically contain from about 30% to about 70% aromatics and are an excellent starting material for obtaining aromatics. The aromatics are purified by solvent extraction process. In general, the reformate is contacted with a lean solvent selective for aromatics in a countercurrent extractor. The aromatics are preferentially dissolved into the solvent phase thereby forming a rich solvent. The rich solvent is withdrawn from one end of the extractor where it is sent to an extractive distillation stripper while the non-absorbed hydrocarbons called the raflinate are withdrawn from the other end of the extractor. Although the stripper is shown in simplified form in the accompanying drawing, the stripper actually perfor-ms two functions, namely the removal of the remaining nonaromatics from the rich solvent phase and the separation of aromatics from the solvent. In order to substantially completely dissolve all the aromatics into the solvent phase in the extractor, Igenerally a small amount of non-aromatics are also dissolved into the solvent phase. The stripper may comprise a series of flash vessels and a fractionator in which case the non-aromatics (and some aromatics) are removed as flashed and distilled overhead vapor, ultimately being recycled to the extractor while the aromatics are removed from the fractionator as a side cut and lean solvent is withdrawn from the bottom of thefractionator. This stripper arrangement is preferably utilized in the case of glycol solvents such as diethyleneglycol, dipropyleneglycol, triethyleneglycol, etc. The stripper may also comprise two fractionators, the first lfractionator being maintained at an intermediate pressure to separate the non-aromatics from the rich solvent by removing them overhead and recycling them to the extractor with the second `fractionator being maintained at a low pressure to separate the aromatics from the solvent by removing the aromatics overhead and the lean solvent out the bottom. This latter stripper arrangement is preferably utilized in the case of solvents such as sulfolane, sulfolene, etc. Other solvents having a selectively for aromatics over more saturated hydrocarbons may be employed in process of this invention such as dimethyl sulfoxide, n-methyl pyrrolidone, etc. with suitable modifications of equipment.

The high purity aromatics are economically more valuable than the raffinate and refiners have attempted to further increase the yield of aromatics by variations in the reforming reactor conditions. For example, lower reactor pressures promote the dehydrogenation of naphthenes and dehydrocyclization of parafiins reactions but lowering the pressure too far results in rapid deactivation of the reforming catalyst. It is of course, impossible to obtain any more aromatics than `is present in the feed and produced in the reforming reactor. I have found that the aromatics present in the feed contribute nothing to the aromatics forming reactions in the reforming reactor and if anything suppress the formation of aromatics. The presence of aromatics in the feed prevent the formation of aromatics on equilibrium limited reactions. This effect is markedly increased in feed wherein a high concentration of aromatics are already present (higher than l5 to 20%) Furthermore, charging aromatics to a reforming reactor permits some ring opening of already existing aromatics to occur. It is accordingly a poor procedure to charge aromatics to a reforming reactor especially when the reformate is subsequently solvent extracted to produce high purity aromatics.

It is an object of this invention to maximize the yield of aromatics from a hydrocarbonaceous feed stock.

Itis another object of this invention to effectively combine an aromatics extraction process and a reforming process to produce enhanced quantities of high purity aromatics from a hydrocarbon feed stock.

It is another object of this invention to minimize the reactor size of a reforming step in an aromatics production process.

It is still another object of this invention to optimize the aromatics forming reaction in a reforming reactor.

It is a specific object of this invention to avoid recycle of a -rafnate which has already once been through 'a reforming reactor.

These and other objects will become more apparent in the light of the following detailed description.

The accompanying figure is a preferable flow scheme for carrying out the process of this invention.

The heart of this invention lies in the extraction of the aromatics from the hydrocarbon feed prior to reforming thereby optimizing the aromatics forming reaction in the reforming reactor and preventing recycle of once reformed raffinate. When the hydrocarbon feed contains at least 15% of aromatics, the pre-extraction step has the following ladvantages: (l) reducing the size if the reforming reactor since liquid hourly space velocity is one of the most significant `factors in determining reactor size and removal of the aromatics reduces the total charge to the reforming reactor; (2) promoting increased dehydrogenation and dehydrocyclization reactions in the reforming reactor since there is 1a greater equilibrium driving force; (3) increasing the hydrogen content of the feed to the reforming reactor which will result in a greater hydrogen evolution per mole of feed. This increased hyd-rogen evolution will give higher nascent hydrogen yield and thereby promote reforming catalyst stability; (4) eliminating any loss of -aromatics already present in the feed; and (5) providing a system wherein once reformed rainate is not further exposed to additional reforming thereby avoiding the disadvantages inherent in trying to further reform refractory raflinate described hereinafter.

The solvent extraction steps are accomplished by introducing the hydrocarbon feed and reformate separately into solvent extractor vessels and therein countercurrently contacting the hydrocarbon phase with the solvent. The countercurrent extraction permits the selective absorption of aromatics into the solvent phase leaving a hydrocarbon phase rich in paraflins and naphthenes. The solvent must have a selectivity for aromatic hydrocarbons over the more saturated parans and naphthenes. Many members of the Iglycol family exhibit this selectivity and are preferably selected from ethylene glycol, diethylene glycol, triethylene glycol, propylene glycol, dipropylene glycol, trimethyene glycol, etc. Another highly preferably family of solvents comprise sulfolane, sulfolene, etc. It is within the scope of this invention to incorporate water with the organic solvent to produce a solvent phase of altered solubility-selectivity relationship. Before the solvent is introduced into the extractor vessel it Eis referred to as a lean solvent since it contains no significant [amounts of dissolved carbons. When the solvent phase is withdrawn from the extractor vessel, it contains an appreciable amount of dissolved hydrocarbons (predominantly aromatics) and is referred to as a rich solvent. The extractor vessel contains suitable contacting means such as sieve decks, packing, rotating discs, etc., in order to promote eflcient contacting `of the solvent phase with the hydrocarbon phase. The solvent phase is essentially immiscible with the hydrocarbon phase but there is a tendency for some hydrocarbons to dissolve into the solvent phase and some solvent to dissolve into the hydrocarbon phase. Whenl one of'the preferable solvents is employed, the hydrocarbons that dissolve in the solvent phase are predominantly aromatic. By selecting the proper solvent to feed ratio and the correct conditions of temperature and pressure, essentially all of the aromatics may be extracted out of the hydrocarbon phase and into the solvent phase with only minor amounts of non-aromatics. Water, which is soluble in the solvent phase, is added to the solvent to increase the selectivity of the overall solvent phase for aromatics although this generally reduces the solubility of the aromatics in the solvent phase. In the case of glycols, water is added in Concentrations of from about 2% to about 20% of the total lean solvent phase while lower amounts are employed in the case of sulfolane.

The rich solvent is sent to an apparatus to remove the remaining non-aromatics, and separate the dissolved aromatics from the solvent phase. This is readily accomplished in a stripper maintained at a lower pressure and higher temperature than the extractor vessel. In the case of a glycol solvent, the non-aromatics are normally removed overhead frorn the stripper along with some of the aromatics and perhaps small amounts of water and solvent whereupon the overhead stream is condensed and separated into an aqueous phase and a hydrocarbon phase. The hydrocarbon phase is generally returned to the lower part of the extractor vessel where these light nonaromatics tend to displace heavier non-aromatics from the rich solvent phase. A side cut stream is removed from the stripper comprising aromatic hydrocarbons without any significant non-aromatic hydrocarbons therein. The glycol solvent is removed from the bottom of the stripper where the water content is adjusted and the now lean solvent is recycled to the extraction vessel. A drag stream of solvent may be continually regenerated by withdrawing a portion of the solvent from the recycle stream, regenerating it and returning it to the recycle stream. Tar-like polycyclic materials tend to build up in the solvent phase and hamper the extraction step. The concentration of these polycyclic materials may be controlled by continually regenerating a drag stream of the solvent. Typically, the solvent is regenerated by introducing the drag stream into a fractionating tower whose operating conditions of pressure and temperature are selected to remove the solvent overhead. The undesirable tar-like materials are removed from the bottom of the fractionating tower.

When utilizing sulfolane as a solvent the stripper is composed of two fractionators, the iirst one being maintained at a pressure lower than the extractor. This first fractionator operates as an extractive distillation column which removes overhead a vapor comprising the remaining nonaromatics present in the rich solvent phase withdrawn from the extractor, some lighter aromatics and perhaps small amounts of sulfolane and water. The overhead vapor is condensed and separated into an aqueous phase and a hydrocarbon phase. The hydrocarbon phase is returned to the lower portion of the extractor to displace heavier non-aromatics from the rich solvent phase. A bottoms stream is withdrawn from the tlirst fractionator and essentially contains aromatic hydrocarbons, sulfolane and some water. This bottoms stream is introduced into a second fractionator maintained at a pressure lower than the first fractionator (preferably under vacuum) wherein the aromatics hydrocarbons are withdrawn overhead and the sulfolane solvent is withdrawn at the bottom. The sulfolane is thereupon recycled to the extractor and in some cases a portion thereof is recycled to the first fractionator to improve the extractive distillation therein.

It is preferable to water wash the raflinate from the hydrocarbonaceous feed extractor (first extractor) prior to reforming to remove remaining traces of solvent from the rainate in order to prevent adverse reaction with the reforming catalyst and to conserve solvent.

The reforming step is accomplished by introducing the ranate from the rst solvent extraction step into a catalytic reactor maintained at reforming conditions in the presence of hydrogen. Typical reforming conditions are temperatures of from about 850 F. to -about 1050" F., pressures of from about 100 p.s.i.g. to about 1000 p.s.i.g., liquid hourly space velocities (LHSV) of from about 0.5 to about 4 and hydrogen to oil mole ratios of from about 2 to about 20. The catalyst may be any catalyst having an appreciable degree of dehydrogenation and cracking activity although I prefer a platinum-halogen catalyst on an alumina support. Preferably the platinum is incorporated in the finished catalyst in concentrations of from about 0.1 wt. percent to about 5 wt. percent and the halogen is selected from the group consisting of chlorine and lluorine. Especially preferably reforming catalysts are platinum and chloride on alumina, the concentration of platinum being from 0.1 to 2.0 wt. percent and the chloride concentration being from 0.1 to 2.0 wt. percent. Since the reforming reactions have the net effect of producing hydrogen, the reactor effluent is separated in a high pressure separator to form a gas phase having an appreciable concentration of hydrogen and a liquid phase. A portion of the gas phase is recycled to the reactor to supply the required hydrogen to oil ratio while the remaining gas is removed from the system as excess net gas. The aromatics forming reactions can be increased by operating at as low a pressure and as high a temperature as practical consistent with catalyst stability considerations. Higher hydrogen purities in the recycle gas and higher hydrogen to oil ratios permit the lowering of pressure while maintaining a constant sta-bility. Since the feed to the reforming reactor has had the aromatics removed, the feed is richer in naphthenes. Naphthenes are readily dehydrogenated in the reforming reactor to form aromatics while simultaneously evolving 3 moles of hydrogen. When paraftins are dehydrocyclized, 4 moles of hydrogen are evolved. Since the aromatics in the feed have already been removed, the equilibrium driving force for the dehydrogenation and the dehydrocyclization reactions is increased thereby not only permitting a greater quantity of aromatics to form but also increasing the hydrogen evolution which may increase the -purity of the recycle gas. This in turn improves catalyst stability which permits increased severity of operating conditions (lower pressure and/or higher temperatures) which in t-urn further increases the yield of aromatics. Therefore when the reformate is introduced into the second extraction vessel, a greater overall quantity of aromatics will be extracted.

The process of this invention has a further advantage, namely, the use of common solvent stripping, solvent regeneration and aromatic fractionators. Since the rich solvents from the two extraction vessels are combined, the increased load is only an incremental loading increase on the fractionation, stripping and regeneration equipment and is considerably less expensive than two separate stripping, regeneration and fractionation systems.

Suitable charge stocks are mixtures of hydrocarbons boiling in the gasoline boiling range, that is those hydrocarbons within the boiling range of hydrocarbons having from 5 to about 10 carbon atoms per molecule. These charge stocks are prepared directly from crude oil by crude lfractionation. Especially preferable charge stocks have from 6 to 9 carbon atoms per molecule. Char-ge stocks prepared from crude oil generally contain paratins, naphthenes and aromatics. The concentration of these types of hydrocarbons vary depending on the source of the crude. Thus California crudes contain an appreciable concentration of aromatics (up to 40%), Midcontinent crudes contain an intermediate concentration of aromatics (up to about 25%) and Kuwait crudes contain a low concentration of aromatics (up to or 15%). Charge stocks which are especially suitable for the process of this invention contain an appreciable amount of aromatics (at least to 20%). Lower concentrations of aromatics will notappreciably improve the reforming reactions as described hereinbefore. l

The process is further described by reference to the accompanying simplified drawing. The naphtha hydrocarbon feed is introduced through flow conduit 1 into extraction vessel 2. Lean solvent is also introduced into extraction vessel 2 through ow conduit 3. The rich solvent phase is withdrawn from the bottom of vessel 2 through flow conduit 21 where it is sent to steps described hereinafter. The ratnate hydrocarbon phase is withdrawn from the top of vessel 2 through ow conduit 4 where it mixes with high pressure separator gas flowing in flow conduit 5 and the mixture flows through ow conduit 6 and into reforming reactor 7. The reforming reactions described hereinbefore are carried out over reforming catalyst 8 and the reactor eluent is withdrawn from reactor 7 through ilow conduit 9. The reactor eluent flows through ow conduit 9 and into high pressure separator 10. Separator gas is withdrawn from separator 10 through llow conduit 11 where a portion flows through flow conduit 12 and out of the system as net separator gas. The remaining recycle gas continues through flow conduit 11 where it passes through compressor 13, through flow conduit 5 and mixes with the raffinate as described hereinbefore. The separator liquid is withdrawn lfrom separator 10 through ow conduit 14 Where it passes into fractionator 15. Since another of the reforming reactions that readily occur is hydrocracking, small amounts of light hydrocarbons such as propane, butane, and pentane are formed in the reactor and are present in the separator liquid. Since these light hydrocarbons contain no aromatics they are preferably removed from the separator liquid before the liquid is solvent extracted. The light hydrocarbons are removed from fractionator 15 overhead through flow conduit 16 and are withdrawn from the system. Some of these light hydrocarbons are useful in gasoline blending.

The stripped separator liquid is withdrawn from stripper 15 through flow conduit 17 where it flows into extractor vessel 18. Lean solvent is also introduced into extraction vessel 18 through flow conduit 29. The rainate hydrocarbon phase is withdra-wn from the top of vessel 18 through flow conduit 19 where it may be utilized in gasoline blending, or for solvent purposes. It is preferred that this entire raffinate not be further recycled to the reforming reactor since it is very refractory and is not readily upgraded by producing further aromatic hydrocarbons.v

This is one of the principal reasons for using two extractors, namely to avoid recycling of this once-reformed raffinate. This raffinate is not readily dehydrocyclized and will only result in increasing the size of the reforming reactor without any real advantage. It is estimated that if only one extractor were used, the size of the reforming reactor would be increased about 35 to 40 volume percent over the present process of this invention. Since if only one extractor were used, the extractor itself would have to be larger, it is expected that the dual extraction process of the present invention is decidedly less expensive and more efficient than a process employing a. cornmon extractor. The rich solvent phase is Withdrawn from the bottom of vessel 1,8 through flow conduit 20 where it combines with the rich solvent flowing in flow conduit 21 and the combined mixture flows through tlow conduit 22 into solvent stripper 23. Stripper 23 is shown as a single fractionator for simplicity but it s intended to employ a proper stripping operation as described hereinbefore depending on the solvent employed. The aromatics are separated from the solvent phase in stripper 23 and are removed throu-gh flow conduit 24. Generally, these aromatics are subsequently fractionated in equipment not shown to produce high purity benzene, toluene, the individual C8 aromatic isomer streams and a Cg-laromatics stream. It is difficult to separate the para-xylene isomer from the meta-xylene by fractionation and preferably freezing techniques are employed. Lean solvent is withdrawn from stripper 23 through iiow conduit 25. A slip stream of lean solvent is lwithdrawn through flow conduit 26 where it is sent to regeneration facilities not shown. The regenerated solvent is returned from the solvent regeneration facilities through. flow conduit 27 where it mixes with the remaining lean solvent and the combined lean solvent ows through flow conduit 28. A portion of the lean solvent flowing in flow conduit 28 is removed through ow conduit 29 where it flows into vessel 18 as described hereinbefore. The remaining portion of the lean solvent flows through flow conduit 3 and into vessel 2 as described hereinbefore.

Auxiliary equipment necessary for the proper functioning of the equipment shown in the figure such as pumps, heat exchangers, heaters, coolers, condensers, control valves, means for actuating control valves, etc. have been omitted in the interest of brevity and simplicity. However, it is to be understood that this auxiliary equipment is necessary for the process to function and it is intended that such equipment `be employed in the process of this invention although its selection is within the ordinary skill of a process and instrumentation engineer.

The following examples are included to further illustrate the operability and usefulness of the present process but it is not intended to limit the invention to the materials presented therein.

EXAMPLE I I into a solvent extraction zone to remove the aromatics and charging the resulting raffinate to a reforming process. The naphtha feed of Example I is introduced into a diethylene glycol solvent extraction contactor at a rate of 10,000 b.p.s.d. and 7500 b.p.s.d. of raffinate and 2500 b.p.s.d. of extract are ultimately produced therefrom. The rainate from the solvent extraction step having a 58 API gravity, an average molecular weight of 105.3, an Engler initial boiling point of- 180 F. and an Engler end point of 320 F. is introduced into a reforming catalyst bed at a rate of 7500 b.p.s.d. The catalyst bed is maintained at temperatures of about 990 F., pressures of about 250 p.s.i.g., a LHSVof 1.5 and a hydrogen to oil mole ratio of 6. Again the reforming process is operated as described hereinbefore producing a net separator gas stream, a light hydrocarbon stream and a liquid reformate stream. The following yields are obtained from the reforming process under these conditions:

Hydrogen standard cu. ft. 7,120,000 Benzene barrels-- 487 Toluene do 562 C8 aromatics do 847 Cg-laromatics do 1778 Total aromatics do 3674 C54- total yield do 6167 Methane pounds 27,600

In addition the F-l clear octane number of the C54- liquid is 90 and the hydrogen p-urity of the recycle gas is 81.6 mole percent. To these yields must be added the extract yields from the solvent extraction step in order to compare the results of Examples I and II. The following Table I shows the results of the over all yield comparison on the basis of 10,000 b.p.s.d. charge stock.

TABLE I Example I, Example II, Example II, Total Rctormate Relormate Extract Example II Yields Benzene, b.p.s.d 824 487 500 987 Toluene, b.p.S.d 085 562 600 l, 162 Cs Aromatics, b.p.s.d 1,110 847 600 1, 447 Cg-l- Aromatics, b.p.s.d 1, 040 1, 778 600 2, 378

Total Aromatics. 4, 859 3, G75 2, 300 5, 974 H2 Yield, s.e.f./d 6. 5x10l 7.12)(100 7.12)(10s (JH-Total Yield, b.p.s.d 8, 606 6, 167 2, 500 8, 667 C5+I `1 Clear Octane No 95 90 08 percent paraftins, 30 volume percent naphthenes and 25 volume percent aromatics is introduced directly into a reforming catalyst Vbed maintained at temperatures of about 980 F., and pressures of labout 250 p.s.i.g. The feed is introduced into the catalyst lbed at a rate of 10,000 barrels per stream day (b.p.s.d.) and the reactor catalyst bed is sized to maintain a LHSV of 2.0 and a hydrogen to oil mole ratio of 6 therein. The reforming process is operated as described hereinbefore and a net separator gas stream, a light hydrocarbon stream and a liquid reformate stream are produced as net products from the process. The following daily yields yare obtained from the operation of the process at these conditions:

Hydrogen standard cu.. ft.-- 6,500,000 Benzene barrels 824 Toluene do 985 C8 aromatics do 1110 C9| aromatics do 1940 Total aromatics do 4859 C5-ltotal yield do 8606 Methane pounds-- 33,000

In addition the F-l clear octane number of the C5-lliquid is 95 and the hydrogen purity of the recycle gas is 80.7 mole percent.

EXAMPLE II This example is presented to show the effect of charging the high aromatic content naphtha feed of Example Comparison of the results of Example I and Example II shows that on the basis of 10,000 b.p.s.d. of charge stock, 163 barrels of additional benzene, 177 barrels of additional toluene, 337 barrels of additional C8 aromatics, 438 lbarrels of additional Cg-iaromatics and 620,000 s.c.f. of additional hydrogen are produced by first solvent extracting the feed followed by reforming the resulting raffinate. In addition, the hydrogen purity of the recycle -gas has increased in Example II (81.6) over Example I (80.7). It is also expected that the reforming catalyst stability in Example I and Example II will be about the same for, although Example II is 10 F. higher in reactor temperature, the recycle hydrogen purity is about 1% higher. The reforming reactor size is exactly the same for both examples although the reactor in Example II could be reduced in size and either operated at more severe conditions or operated at the same conditions and produce a lower aromatics yield. Since Example II produces a total aromatics yield increase of 11.1 volume percent over Example I, if the charge stock is solvent extracted and the raffinate is reformed at the same conditions as in Example I (at 2.0 LHSV) there results an increase in total aromatics yield in excess of 5 volume percent and a reduction in the reforming reactor size of 25%. Likewise other operating conditions in the reforming reactor may be selected to optimize the economic advantage of solvent extracting a charge stock yand reforming the resulting ranate.

I claim as my invention:

1. A- process for the enhanced productionY of aromatic hydrocarbons from a hydrocarbonaceous feed containing aromatics which comprises the steps:

(a) solvent extracting the feed in a first extractor to produce a first rich solvent phase containing aromatic hydrocarbons and fa first hydrocarbonaceous rainate phase;

(b) reforming the first raffinate phase by contacting the raffinate with a reforming catalyst in a reaction zone maintained at reforming conditions;

(c) solvent extracting the reformed raffinate in a second extractor to producel a second rich solvent phase containing aromatics and a second hydrocarbonaceous raffinate phase;

(d) separating the aromatics from the first and the second rich solvent phases; and

(e) recovering the laromatics and said second rainate phase.

2. The process of claim 1 further characterized in that the first and second rich solvent phases are combined, and stripped to remove non-aromatics from the phase prior to separating the aromatics from the solvent.

3. The process of claim 1 further characterized in that the hydrooarbonaceous feed contains at least l5 volume percent aromatic hydrocarbons.

4. The process of claim 3 further characterized in that the solvent employed in the solvent extraction steps comprises sulfolane.

5. The process of claim 3 further characterized in that the solvent employed in the solvent extraction steps comprises at least one glycol selected from the group consisting of diethyleneglycol, triethyleneglycol and dipropyleneglycol.

6. The process of-claim 3 further characterized in that the reforming catalyst comprises an alumina support having both platinum and at least one halogen selected from chlorine and uorine thereon and the reforming conditions comprise temperatures of from about 850 F. to about 1050 F., pressures of from about 100 p.s.fi.g. to about 1000 p.s.i.g., liquid hourly space velocities of from about 0.5 to about 4 and hydrogen to oil mole ratios of from about 2 to about 20.

7. The lprocess of claim 1 further characterized in that the first -rafiinate phase is water washed prior to the reforming step.

8. The process of claim 1 further characterized in thlat the normally liquid reformed raffinate is fractionated to remove C5 and lighter hydrocarbons prior to the second solvent extraction step.

References Cited UNITED STATES PATENTS 2,915,453 12/1959 Haensel et al. 208-64 2,976,231 3/ 1961 Bloch 208-96 2,981,675 4/1961 Hemminger et al 208-95 2,838,582 6/ 1958 Kassel et al 208-87 2,972,646 2/1961 Cahn et al 260-68l.5 3,008,895 11/1961 Hansford et al. 208-68 3,222,416 12/ 1965 Evans et al 260-674 DELBERT E. GANTZ, Primary Examiner.

G. E. SCHMITKONS, Assistant Examiner. 

